Process for control of polymer fines in a gas-phase polymerization

ABSTRACT

A gas phase polymerization process comprising: ( 1 ) preparing a solution of a catalyst precursor comprising a mixture of magnesium and titanium compounds, an electron donor and a solvent; ( 2 ) adding a filler to the solution from step ( 1 ) to form a slurry; ( 3 ) spray drying the slurry from step ( 2 ) at a temperature of 100 to 140° C. to form a spray dried precursor, ( 4 ) slurring the spray dried precursor from step ( 3 ) in mineral oil, ( 5 ) partially or fully pre-activating the catalyst precursor by contacting the slurry of ( 4 ) with one or more Lewis Acids, and ( 6 ) transferring the partially or fully activated precursor from step ( 5 ) into a gas phase reactor in which an olefin polymerization reaction is in progress.

CROSS REFERENCE TO RELATED APPLICATIONS

This application claims the benefit of U.S. Provisional Applications60/469,663 and 60/469,665, both filed May 12, 2003.

BACKGROUND OF THE INVENTION

The use of Ziegler-Natta catalyst systems to promote various olefinpolymerizations is well known. These catalyst systems function in boththe gas phase, and slurry as well as solution polymerization processes.Of these processes, the gas phase and slurry polymerization processesare also known as particle form processes, socalled because the polymeris formed as discrete particles, the size and shape of which is afunction of the size and shape of the catalyst particle. The polymerparticle is thus said to replicate the initial catalyst particle. Thefinal size of the polymer particle is a function of both the initialcatalyst particle size and the productivity of the catalyst. Thus, inpreparing a catalyst to be used in a gas-phase polymerization process,great care is taken in the catalyst precursor preparation in order tocontrol both polymer particle size and morphology in addition toproductivity. Examples of such Ziegler-Natta catalysts include thosedisclosed in U.S. Pat. Nos. 4,302,565; 4,482,687; 4,508,842; 4,990,479;5,122,494; 5,290,745; and, 6,187,866.

Another polymer property that is desirably controlled through catalystcontrol is the particle size distribution, especially with respect tothe low end of the distribution, as an unacceptable amount of smallcatalyst particles could lead to the generation of small polymerparticles known as “polymer fines”. Polymer fines are undesirable in gasphase fluidized bed polymerization systems, as they tend to segregate tothe top of the fluidizing bed, causing problems with bed level control.They are also preferentially entrained into the cycle gas leading tosystem plugging in heat exchangers and compressors, buildup in thebottom head of the reaction system and formation of gels due tocontinued polymerization at lower temperatures than the bulk of thepolymer product. All of the above lead to poor commercial operation,reduced polymerization efficiency, and generally impaired operation.High levels of fines can also cause problems in downstream handling ofthe polymer once it exits the polymerization system. Fines can causepoor flow in purge bins, plug filters in bins and present safetyproblems. The above problems make elimination or reduction of polymerfines important to commercial operation of a gas-phase polymerizationprocess.

In a multiple series reactor system, where the composition of thepolymers produced in the separate reactors is often widely variant, thepresence of polymer fines is particularly harmful to continuous andsmooth operation. This is due to the extreme importance of precise bedlevel control as the product properties of the polymer are stronglyinfluenced by the relative amount of polymer produced in each reactor.If the bed weights are not precisely known, it is extremely difficult toproperly control the product exiting the final reactor.

With respect to the preparation of linear low density polyethylene andother ethylene/α-olefin copolymers, it is preferred to produce polymerin the separate reactors with both large molecular weight differencesand relatively large differences in incorporated comonomer. To producefinal polymers with the best physical properties, it is preferred tohave one of the reactors produce a polymer with high molecular weightand incorporating a majority of the comonomer. In the second reactor, alow molecular weight portion of the polymer is formed which may alsohave comonomer incorporated, but normally in an amount less than thatincorporated in the high molecular weight portion. When the highmolecular weight component is produced first, polymer fines can become asignificant problem, especially when the flow index (I21) of theresulting polymer is in the range from 0.1 to 2.0 g/10 min, and theincorporated comonomer content is less than 5 weight percent, especiallyless than 4.5 wt weight percent.

Depending on the order of production of the different polymers in themultiple reactor system (that is high molecular weight first, lowermolecular weight second or the reverse), the fines will tend to havesignificantly different polymer properties than the bulk of the polymergranules. This is due to the fact that the fines also tend to be theyoungest particles in the reactor and hence have had insufficientresidence time in the reactor to produce a representative amount ofpolymer before transit to the second reactor in series.

This in turn leads to further problems in compounding of the polymerinto pellets for end-use. In particular, the fines are normally ofsignificantly different molecular weight or branching compositioncompared to the bulk polymer. Although the particles of both the bulkmaterial and the fines will melt at roughly the same temperature, mixingis hampered unless the products have a similar isoviscous temperature(that is the temperature at which the melt viscosity of the two productsis essentially the same). These polymer fines, which tend to be ofsignificantly different molecular weight than the bulk of the polymerand differing isoviscous temperature, are then poorly mixed with thebulk phase. Upon cooling after pellet formation, these poorly mixedregions, if of sufficient size, will be visible in blown films as gelsor in other extruded articles, resulting in visual defects and stressconcentrators leading to premature failure of an article made therefrom.

Thus, polymer fines are, in general a problem for gas phase olefinpolymerization processes and, in particular an issue for staged orseries reactor systems in which precise control of polymer compositionis only achieved by precise control of the relative amount of polymerproduced in the multiple reactors.

Polymer fines can be removed from the polymerization reactor though useof a cyclone on the recycle line, however this reduces productivity andincreases operating costs. In addition, the fines tend to be higher incatalyst concentration as they are, on average, younger particles.Removing these particles from the polymerization reactor increases theneed for fresh catalyst further increasing costs. Since the polymerfines are still active for further polymerization, special care must betaken to make sure that the fines do not plug the cyclone. Any areas inwhich polymer particles can congregate in the presence of olefin canresult in continued polymerization leading to formation of agglomeratedparticles and large chunks of polymer.

U.S. Pat. No. 5,969,061 disclosed the use of a solvent in an attempt toreduce polymer fines by making the bulk of the polymer particlesstickier, resulting in the fines attaching to the larger particles.However, increasing polymer stickiness can result in further problemsdownstream in product separation and makes the reaction system morevulnerable to loss of recycle flow due to power failures, increasing therisk of large agglomerate formation. The addition of large amounts ofsolvent also increases the cost and complexity of the reaction systemand requires apparatus for recycle of the solvent for reuse. It would bedesirable to produce fewer fines during the polymerization reaction,thereby reducing the need for other polymer fines control systems.

In gas phase polymerization systems, it is known that, generally, eachcatalyst particle produces one polymer particle. Catalyst particles, ingeneral increase in particle size proportionally to the cube root of thecatalyst productivity. That is, the polymer particle size is expressedby the formula: polymer particle size=Constant×(CatalystProductivity)^(1/3).

While not being bound by any one theory, it is believed that polymerfines originate either from fines in the catalyst or by particleattrition of the growing polymer. Given that fines can still be presentin a polymer produced in the first reactor of a multiple reactorconfiguration producing tough, high mechanical strength, high molecularweight polymer, it is unlikely that particle attrition is the primarycause of polymer fines in such systems. Thus, catalyst particle finesare believed to be the predominant cause of polymer fines. Such catalystfines can be removed by a variety of methods, ranging from eluting tosieving of the catalyst prior to use. This, however, adds both cost andcomplexity to the catalyst preparation process as well as increases thelikelihood of catalyst contamination during the additional processingsteps.

Operating the reaction system at higher levels of catalyst productivitycan also reduce polymer fines. For single reactor systems, this isusually a feasible approach, however operating at catalyst productivitylevels that are too high can result in operability problems due topolymer particle agglomeration. In extreme cases, higher levels of finesdue to fracture of catalyst particles during polymerization may alsoresult. For multiple reactor systems in which the catalyst is added onlyto the first reactor in the series, increasing catalyst productivity inthe first reactor to minimize fines can result in the inability to runthe second (or additional) reactors at commercially feasible conditionsdue to catalyst deactivation.

In order to compensate for this activity loss due to catalystdeactivation, the first reactor of the multiple reactor system is oftenoperated in a “low productivity” regime so that there is sufficientcatalyst activity remaining to complete polymerization in the second(and subsequent) reactors. However, the operation at lower catalystproductivity in the first reactor results in a reduction in polymerparticle size, further increasing the need to control and reduce fineswhich might be caused by the catalyst.

As already explained, the particle size of a given polymer particle is afunction of both the initial catalyst particle size, and theproductivity of the catalyst; that is how much polymer grows from theinitial catalyst particle during the polymerization process. Thus smallparticle sized polymer or polymer fines can be a result of either smallinitial catalyst particle size or low catalyst productivity, or both,and when both conditions are present, the generation of polymer fines isexacerbated.

The catalysts used in many olefin polymerization processes are of theZiegler-Natta type. In particular, for gas phase polymerizations, thecatalyst is often made from a precursor comprising magnesium andtransition metal halides, particularly titanium chlorides in an electrondonor solvent. This solution is often either deposited on a porouscatalyst support, or a filler is added, which, on subsequent spraydrying, provides additional mechanical strength to the particles. Thesolid particles from either method of production are often slurried in adiluent to produce a high viscosity mixture, which is then used in agas-phase polymerization. Exemplary catalyst compositions are describedin U.S. Pat. Nos. 6,187,866 and 5,290,745. Precipitated/crystallizedcatalyst compositions such as those described in U.S. Pat. Nos.6,511,935 and 6,248,831, may also be used. Additional techniques forforming suitable catalyst precursors for use herein are disclosed inU.S. Pat. Nos.: 5,247,032, 5,247,031, 5,229,342, 5,153,158, 5,151,399,5,146,028, 5,124,298, 5,106,806, 5,082,907, 5,077,357, 5,066,738,5,066,737, 5,034,361, 5,028,671, 4,990,479, 4,927,797, 4,829,037,4,816,433, 4,728,705, 4,548,915, 4,547,476, 4,540,679, 4,535,068,4,472,521, 4,460,701, 4,442,276, and 4,330,649.

One advantage of the use of a spray drying process is that it allows theparticle size and morphology of the catalyst, and hence the finalproduct, to be controlled by variation of the process parameters of thespray dryer. Such parameters include the speed of the atomizer, thesolids content of the slurry to be dried, the inlet and outlet gastemperatures of the dryer and the feed rate of the slurry to theatomizer.

However, due to the nature of spray drying, some small particles arealways present. In particular, some “micro-fine” particles are formedduring the spray drying process. These are also frequently called“daughter” particles and result from break up of droplets during thespray drying operation. These particles end up in the final spray driedcatalyst composition and are of essentially the same chemicalcomposition as the larger size, desired particles. These particles areseen in the <10 micron fraction of the particle size distribution of thecatalyst and can form fine polymer particles that are the root ofoperational problems.

The catalyst precursor as produced is essentially inactive for olefinpolymerization due to the presence of the Lewis Base electron donor.Activation of the catalyst precursor requires the removal of theelectron donor from the vicinity of the active site, that is, the metaland, if necessary, reduction of the metal. The activator extracts theelectron donor compound from the active site in one of several ways. Theelectron donor can be removed by complex formation, or by alkylation orby reduction and alkylation if the valence state of the metal requiresreduction. Typical activating compounds are Lewis Acids. The activatoris used to remove at least 90 percent, preferably all or as near to allas possible, of the electron donor from the active site, that is, thetransition metal.

If the Lewis Acid is a non-reducing compound, such as BCl₃, AlCl₃, orsimilar chlorinating agent, a reducing compound, typically atrialkylaluminum an aluminum dialllyl halide may be added to fullyactivate the catalyst precursor. Precursors that are not fullyhalogenated will also require either use of a halogen donating LewisAcid or a separate halogenation step prior to use.

Activation of the catalyst typically occurs in the polymerizationreactor by the cocatalyst. However, complete activation of the catalystinside the polymerization reactor typically requires a substantialexcess of activator compound and in-the case of higher (C₃, C₄ and up)olefin polymerizations, reintroduction of a Lewis base as a selectivitycontrol agent. Advantages to this technique are its simplicity ofcatalyst manufacture and feed. However, use of excess activator compoundnot only leads to added operational expense, but it may causeoperational problems or detriment to the final product. Ultimately,large quantities of activator are required due to dilution by monomers,diluents, condensing agents, and other components within the reactor.

Partial pre-activation can occur prior to the polymerization reactor,however this additional step, because it potentially increases theexposure of the active catalyst to impurities and other deactivators,can cause a decrease in catalyst productivity especially on extendedstorage prior to use. Such a deactivation and loss of productivity wouldin turn be expected to cause an increase in polymer fines. Thus, as fullactivation of the catalyst is normally completed in the polymerizationreactor with excess co-catalyst and generally occurs whether thecatalyst has been partially activated or not, until now there has beenno driving force to pursue partial pre-activation of the catalyst priorto addition to the reactor.

However, it would still be highly advantageous to have a process thatwould minimize the generation of polymer fines in a gas phasepolymerization. It would also be advantageous if this process were to beapplicable to a gas phase process utilizing multiple reactors. It wouldbe even more advantageous if such a process involved a relatively simplemanipulation of the catalyst rather than the more expensive anddifficult process modifications such as cyclone operation or addition ofsolvents to the reactor. Finally, a process in which fully activatedcatalyst composition is supplied to the reactor would additionally bedesirable.

SUMMARY OF THE INVENTION

The present invention is a process for reducing the amount of polymerfines in a gas phase particle form polymerization process by at leastpartially pre-activating the catalyst precursor by the addition of aLewis acid to the catalyst precursor prior to its introduction to thepolymerization reactor.

More specifically, the present invention relates to a gas phase olefinpolymerization process comprising:

-   -   (1) preparing a solution of a catalyst precursor comprising a        mixture of magnesium and titanium compounds, an electron donor,        and optionally a solvent;    -   (2) adding a filler to the solution from step (1) to form a        slurry;    -   (3) spray drying the slurry from step (2) at a temperature of        100 to 140° C. to form a spray dried precursor;    -   (4) slurrying the spray dried precursor from step (3) in mineral        oil,    -   (5) partially or fully pre-activating the catalyst precursor by        contacting the slurry of (4) with one or more Lewis Acids, and    -   (6) transferring the partially or fully activated precursor from        step (5) into a gas phase reactor in which an olefin        polymerization reaction is in progress.        Alternatively the process may comprise:    -   (1) preparing a solution of a catalyst precursor comprising a        mixture of magnesium and transition metal compounds an electron        donor and optionally a solvent;    -   (2) adding a porous catalyst support, to the solution from        step (1) to form a slurry;    -   (3) drying the slurry from step (2) to form a solid catalyst        precursor;    -   (4) re-slurying the solid catalyst precursor from step (3) in        mineral oil,    -   (5) partially or fully activating the catalyst precursor by        contacting the slurry of (4) with a Lewis Acid; and

(6) transferring the partially or fully activated catalyst precursorfrom step (5) into a gas phase reactor in which an olefin polymerizationreaction is in progress.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 shows a schematic flow chart of the catalyst activation processof the present invention.

DETAILED DESCRIPTION OF THE INVENTION

For purposes of United States Patent practice, the contents of anypatent, patent application, or publication referenced herein are herebyincorporated by reference in their entirety (or the equivalent USversion thereof is so incorporated by reference) especially with respectto the disclosure of synthetic techniques, raw materials, and generalknowledge in the art.

If appearing herein, the term “comprising” and derivatives thereof isnot intended to exclude the presence of any additional component, stepor procedure, whether or not the same is disclosed herein. In order toavoid any doubt, all compositions claimed herein through use of the term“comprising” may include any additional additive, adjuvant, or compound,unless stated to the contrary. In contrast, the term, “consistingessentially of” if appearing herein, excludes from the scope of anysucceeding recitation any other component, step or procedure, exceptingthose that are not essential to operability. The term “consisting of”,if used, excludes any component, step or procedure not specificallydelineated or listed. The term “or”, unless stated otherwise, refers tothe listed members individually as well as in any combination.

The term “polymer fines” as used herein means polymer particles of lessthan 125 μm in particle size.

The term “catalyst precursor” as used herein means a mixture comprisingtransition metal and magnesium compounds and a Lewis Base electrondonor. Preferably the catalyst precursor has the formulaMg_(d)(M)(OR)_(e)X_(f)(ED)_(g) wherein R is an aliphatic or aromatichydrocarbon radical having 1 to 14 carbon atoms or COR′ wherein R′ is aaliphatic or aromatic hydrocarbon radical having 1 to 14 carbon atoms;each OR group is the same or different; M is a transition metal,preferably titanium, zirconium, haflium, vanadium or a mixture thereof;X is independently chlorine, bromine or iodine; ED is an electron donor;d is 0.5 to 56; e is 0, 1, or 2; f is 1 to 116; and g is >2 and up to1.5(d)+3. It is prepared by combining one or more transition metalcompounds, a magnesium compound, and an electron donor, optionally in asolvent, and forming a solid particulated product therefrom.

Preferred transition metal compounds are titanium compounds, mostpreferably of the formula Ti(OR)_(e)X_(h) wherein R, X, and e are asdefined above; h is an integer from 1 to 4; and e+h is 3 or 4. Somespecific examples of suitable titanium compounds are: TiCl₃, TiCI₄,Ti(OC₂H₅)₂Br₂, Ti(OC6H₅)Cl₃, Ti(OCOCH₃)₂Cl₃, Ti(acetylacetonate)₂Cl₂,TiCl₃(acetylacetonate), and TiBr₄. TiCl₃ and TiCl4 are preferredtitanium compounds. The magnesium compounds include magnesium halidessuch as MgCl₂, MgBr₂, and MgI₂. Anhydrous MgCl₂ is a preferred compound.Desirably 0.5 to 56, and preferably 1 to 10, moles of the magnesiumcompound are used per mole of titanium compound in forming theprecursor. Vanadium, hafnium and zirconium compounds may be used inadmixture with the titanium component if desired. Specific vanadiumcompounds which may be used are VCl₃, VOCl₃, V(acetylacetonate)₃.Specific zirconium compounds that are useful are ZrCl,, ZrBr₄,ZrCl₂(acetylacetonate)₂, and Zr(OR¹)₄ where R¹=ethyl, n-butyl, isobutyl,n-hexyl or n-octyl. Specific hafnium compounds useful in the inventionare HfCl₄ and Hf(OR¹)₄, wherein R¹ is as previously defined.

Suitable catalyst precursors and methods of producing the same are knownin the art and disclosed for example in U.S. Pat. Nos. 5,034,361;5,082,907; 5,151,399; 5,229,342; 5,106,806; 5,146,028; 5,066,737;5,077,357; 4,442,276; 4,540,679; 4,547,476; 4,460,701; 4,816,433;4,829,037; 4,927,797; 4,990,479; 5,066,738; 5,028,671; 5,153,158;5,247,031; and 5,247,032.

The electron donor is a Lewis base, preferably one that is liquid attemperatures in the temperature range from 0 to 200° C. and in which themagnesium and transition metal compounds are soluble. Examples includealkyl esters of an aliphatic or aromatic mono- or dicarboxylic acids,aliphatic ketones, aliphatic amines, aliphatic alcohols, alkyl-,cycloalkyl-, aryl-, and alkyaryl-ethers, compounds containing mixturesof the foregoing functionality, and mixtures thereof, each electrondonor having 2 to 20 carbon atoms. Preferred are aliphatic andcycloaliphatic ethers having 2 to 20 carbon atoms; dialkyl-, diaryl-,and dialkaryl-ketones having 3 to 20 carbon atoms; dialkyl carbonates,alkylene carbonates, and alkyl-, alkoxyalkyl-, and aryl-esters ofaliphatic or aromatic mono- or dicarboxylic acids or alkoxy-substitutedderivatives thereof, having 2 to 20 carbon atoms. Specific examples ofsuitable electron donors are methylformate, ethylacetate, butylacetate,diethyl ether, tetrahydrofuran, dioxane, di-n-propyl ether, di-n-butylether, ethanol, 1-butanol, ethylformate, methylacetate, ethyl benzoate,ethyl p-methoxybenzoate, ethyl p-ethoxybenzoate, diethylphthalate,diisobutylphthalate, di-n-butylphthalate, diisooctylphthalate, ethylenecarbonate, and ethylpropionate. The most preferred electron donor istetrahydrofuran. Mixtures of electron donors may be used as well.Bicomponet mixtures, that is those mixtures employing Electron Donor (1)and Electron Donor (2)where the mole ratio of Electron Donor(1)/Electron Donor (2) ranges from 0.01:1 to 10:1 with a preferred rangeof 0.01:1 to 1:1. Highly preferably Electron Donor (2) istetrahydrofuran, and it is present in excess. Especially preferredcombinations of electron donor compounds are: ethanol withtetrahydrofuran; 1-butanol with tetrahydrofurn isopropanol withtetrahydrofiran; ethylbenzoate with tetrrhydrofuran, anddiisobutylphthalate with tetrahydrofuran.

While a large excess of electron donor may be used initially to providethe reaction product of titanium compound and electron donor, the finalcatalyst precursor contains 1 to 20 moles and preferably 1 to 10 molesof electron donor per mole of titanium compound. Excess electron donormay be removed by extraction, washing or devolatilization and preferablyis removed by drying during a spray drying process.

Since the catalyst will act as a template for the growth of the polymer,it is essential that the catalyst precursor be converted into a solid.It is also essential that the resultant solid has the appropriateparticle size and shape to produce polymer particles with relativelynarrow size distribution, low amounts of fines and good fluidizationcharacteristics. Although this solution of Lewis Base, magnesium andtransition metal compounds may be impregnated into a porous support anddried to form a solid catalyst, it is preferred that the solution iscombined with a filler and converted into a solid catalyst via spraydrying.

Formation of the Catalyst Precursor

The catalyst precursor may be prepared according to any suitabletechnique for preparing a particulated, solid product containing theelectron donor. A preferred catalyst precursor comprises TiCl₃, formedby the reduction of TiCl₄ with magnesium metal in solution in theelectron donor solvent. The electron donor employed in this embodimentof the invention must be free of substituents containing activehydrogen, such as hydroxyl groups, as such functional groups readilyreact with both magnesium and titanium tetrachloride. Reduction oftitanium tetrachloride with magnesium metal according to the presentinvention takes place in a solvent comprising the electron donor andresults in the formation of magnesium dichloride and titaniumtrichloride, which then form complexes with the electron donor. Thisreaction can be illustrated by the following equation: 2 TiCl₄(ED)₂+Mg→2TiCl₃(ED)₃+MgCl₂(ED)_(1.5)

where ED is as previously defined, preferably tetrahydrofuran.

Because magnesium is highly reactive with titanium tetrachloride, it ispreferable to employ the metal in granular form rather than as a powder.The use of larger granular particles of the metal rather than the moreminute powder form limits the reactivity of the metal and allows thereaction to proceed in a smoother and more controlled manner. Proceedingin this manner also limits over-reduction of titanium tetrachloride totitanium dichloride, which might otherwise occur. Usually magnesiumparticles having an average particle size of from 0.25 mm to 10 mmpreferably from 1 mm to 4 mm, are employed.

Reduction of titanium tetrachloride to titanium trichloride is effectedusing an essentially stoichiometric amount of the magnesium metalrequired to effect the reduction, that is, one mole of magnesium metalis employed for every two moles of titanium tetrachloride. At least astoichiometric amount of magnesium is required to completely reduce thetitanium tetrachloride to titanium trichloride. On the other hand, anexcess of magnesium is undesirable as such excess must then be removedfrom the reaction mixture. In addition, use of excess magnesium cancause over-reduction of titanium tetrachloride to titanium dichloride.

From 5 mols to 400 mols of electron donor compound are advantageouslyemployed per mol of titanium tetrachloride, preferably 50 mols to 200mols of electron donor compound per mol of titanium tetrachloride, withmost of the residue being removed as explained earlier.

Usually the magnesium metal is added to a mixture of titaniumtetrachloride dissolved in the electron donor compound. However, it isalso possible to add the titanium tetrachloride to a mixture of themagnesium metal in the electron donor compound, or even to add thetitanium tetrachloride and magnesium metal to the electron donorcompound together. Ordinarily reaction is effected below the boilingpoint of the electron donor compound, preferably between 20 and 70° C.An inert atmosphere should be maintained, that is, an atmosphere that isnonreactive under the conditions employed during the reduction.

The reduction of titanium tetrachloride with magnesium metal results ina solution, which contains one mole of magnesium dichloride for everytwo moles of titanium trichloride, and which is substantially free ofundesirable by-products. In order to complete formation of the desiredcatalyst precursor, it is only necessary to add additional magnesiumdichloride to the solution to increase the Mg/Ti ratio to the desiredlevel. The solution can then be impregnated in a suitable support, orspray dried with or without a suitable filler, to obtain discreteparticles of the desired precursor.

After the completion of the magnesium metal reduction of the TiCl₄,additional transition metal compounds such as those defined previouslymay be added. Additional electron donor compounds, especially thosewhich may have reactive functionality towards either Mg metal or TiCl₄may be added as well. When added, the ratio of second transition metalcompound to the Ti will range from 0.1:1 to 10:1 and preferably 1:1 to5:1.

Magnesium dichloride may be added to the solution to increase theMg/transition metal ratio depending upon whether the solution is to beimpregnated in a suitable support or spray dried. Because drying is notconstrained to occur completely within the pores of a support when thesolution is spray dried, higher amounts of magnesium dichloride areordinarily employed when this procedure is followed than when thesolution is impregnated in a support. Generally, sufficient magnesiumdichloride is added to the solution to increase the Mg/Ti ratio to arange of from 1:1 to 56:1, preferably to a range of from 1.5:1 to 5:1.When the solution is to be spray dried, it is preferable to addsufficient magnesium dichloride to increase the Mg/Ti ratio to a rangeof from 1.5:1 to 15:1, most preferably to a range of from 4:1 to 6:1.

Dissolution of magnesium dichloride can be effected by stirring it inthe electron donor solution at a temperature of from 0 to 200° C.Temperatures that are hotter than the boiling point of the electrondonor compound may be utilized, however this requires the use ofequipment capable of withstanding elevated pressures, and for thisreason is generally not preferred. Magnesium dichloride more readilydissolves in the electron donor compound in the presence of titaniumtetrachloride than in the presence of titanium trichloride. Thus, inthose instances when the titanium tetrachloride is to be reduced totitanium trichloride by adding the magnesium metal to a solution of thetitanium tetrachloride in the electron donor compound, it may bepreferable to dissolve both the magnesium dichloride and the titaniumtetrachloride in the electron donor compound before the magnesium metalis added. The magnesium dichloride can, of course, also be dissolved ina mixture of the magnesium metal and electron donor compound before thetitanium tetrachloride is added to the mixture, if desired.

The solution of titanium trichloride and magnesium dichloride preparedin this manner can be spray dried as is, however the particles thusformed are typically brittle and exhibit insufficient mechanicalstrength leading to increased levels of fines. It is thought that theseparticles are relatively brittle due to the highly crystalline nature ofthe solids formed. It is preferable to either deposit the precursorsolution on a porous catalyst support or instead add to the solution afiller which, on subsequent spray drying, provides additional mechanicalstrength to the particles.

When the precursor solution is deposited on a porous support, thesupport chosen is inert, that is, it does not affect the polymerizationreaction in itself. However, when the precursor is deposited on thesurface of a support having a large surface, the monomer molecules aremore readily polymerized. The support is either an organic compound (forexample a polymer) or an inorganic compound, such as a metal oxide.Suitable inorganic compounds include, for example, silicon dioxide,aluminum oxide. Ti-, Mg-, Cr-, Ba-, Th- and Zr-oxides, silicates,aluminophophates, and mixtures of alurina and aluminum phosphate(phosphated alumina). The inorganic support can also be a metalhydroxide or a metal hydroxy halide. Combinations of various supportsare possible, as well. The amount of support used to form the catalystprecursor ranges from 50 to 90, preferably from 70 to 85 percent of thetotal catalyst precursor. The support should also be chosen such thatthe precursor solution prior to drying is contained substantiallyentirely within the pores of the support and is deposited therein byprecipitation during the drying step.

Alternatively, and more preferably, spray drying may be effected by i)admixing the precursor solution with said filler; ii) optionally heatingthe resulting slurry to a temperature as high as the boiling point ofthe electron donor compound; and iii) then atomizing the slurry by meansof a suitable atomizing device to form discrete spherically shapedparticles. Atomization is effected by passing the slurry through theatomizer together with an inert drying gas, that is, a gas that isnon-reactive under the conditions employed during atomization. Anatomizing nozzle or a centrifugal high-speed rotary atomizer can beemployed to effect atomization. The volumetric flow of drying gas mustconsiderably exceed the volumetric flow of the slurry to effectatomization of the slurry and removal of excess electron donor compound.Ordinarily the drying gas is heated to a temperature greater than theboiling point of the electron donor compound up to as high as 200° C. tofacilitate removal of excess electron donor compound. However, if thevolumetric flow of drying gas is maintained at a very high level or ifreduced pressures in the spray drying apparatus are employed, lowertemperatures may be used.

Any solid particulate material that is inert to the other components ofthe catalyst system, and during subsequent polymerization, can beemployed as filler for the solution of titanium trichloride andmagnesium dichloride to form a slurry suitable for spray drying. Suchmaterials can be organic or inorganic. Suitable fillers include silica,titanium dioxide, alumina, aluminophosphates, talc, polystyrene, andcalcium carbonate. Fumed hydrophobic silica is preferred because itimparts high viscosity to the feed slurry, is inert in the finalproduct, and provides good strength to the spray dried particles. Anexample of such a hydrophobic fumed silica includes Cab-O-Sil™,available from the Cabot Corporation.

The particulate material employed as filler should have an averageparticle size no greater than 10 μm, preferably no greater than 1 μm.Like the particulate materials employed when the solution of titaniumtrichloride and magnesium dichloride is impregnated into a support, theparticulate material employed as filler should be substantially free ofabsorbed water and unreactive with the remaining catalyst components.Filler compounds which are soluble in the electron donor solvent mayalso be used. Examples include CaCl₂, polyvinylchloride, polystyrene,interpolymers of styrene and ethylene, and acrylic polymers. Soluble andinsoluble fillers may be used separately or in mixture. When used in amixture, the weight ratio of soluble filler:insoluble filler ispreferably from 0.05:1 to 1:1.

When an insoluble filler is used, sufficient filler should be admixedwith the solution of titanium trichloride and magnesium dichloride toproduce a slurry suitable for spray drying, that is, a slurry containingsuch filler in an amount of from 0 to 15, preferably from 2.5 to 10percent by weight. When spray dried, such slurry produces discreteparticles in which filler is present in an amount of from 0 to 50,preferably from 10 to 50 percent by weight, most preferably 15 to 30percent by weight. The spray dried particles desirably have an averageparticle size of from 5 to 200 μm, preferably from 15 to 50 μm

Spray drying is effected using any suitable apparatus known in the art.Due to the particle size desired, rotary atomization is the preferredmethod to convert the feed slurry into droplets for drying. A co-currentdrying chamber is preferably employed in which the aspect ratio (H/D) isbetween 0.8 and 3.0, preferably near 1.0. A closed cycle spray dryersystem is also preferred for use if flammable electron donors or othercomponents are employed.

The spray dried catalyst precursor is then preferentially placed intomineral oil slurry. The mineral oil used for the formation of the slurrycan be any essentially air and moisture free aliphatic or aromatichydrocarbon, preferably an aliphatic hydrocarbon, which is unreactivewith the catalyst precursor composition, the activator, and thecocatalyst. Suitable diluents include hydrogenated mineral oils,including aliphatic or naphthenic oils of relatively high viscosity tominimize settling of catalyst solids in feed tubes, although, withappropriate engineering design, lower viscosity diluents such asisopentane, hexane, and heptane can be used as well. Preferred diluentsare aliphatic or napthenic hydrocarbons with viscosity greater than 50centipoise (cP) particularly greater than 70 cP and less than 5,000 cP,as measured by a Brookfield viscometer at a shear rate of 1 sec⁻¹ at 25°C. The viscosity of the diluent is sufficiently low so that the slurrycan be conveniently pumped through the pre-activation apparatus andeventually into the polymerization reactor, using a slurry catalystfeeder. Progressive cavity pumps for large volume flows and dual pistonsyringe pumps, where the catalyst flows are ≦10 cm³/hour of slurry, aresuitably employed. Particularly preferred diluents are food grademineral oils, exemplified by Kaydo™ 350 and Hydrobrite™ 380, 550 and1000, available from Witco Corporation.

Pre-Activation of the Catalyst Precursor

Prior to its introduction into the reactor, the catalyst precursor iscontacted with a Lewis acid activator. The Lewis acid activator can beone compound or a mixture of two or more different compounds. PreferredLewis acids are those of the formula M′(R″_(n))X_((3-n)) wherein M′ isaluminum or boron; each X is independently chlorine, bromine, or iodine;each R″ is independently a saturated aliphatic hydrocarbon radicalhaving 1 to 14 carbon atoms, provided that when M′ is aluminur n is 1 to3 and when M′ is boron, n is 0 to 1.5. Examples of suitable R″ groupsare methyl, ethyl, n-butyl, isobutyl, n-hexyl, n octyl, n-decyl, andn-dodecyl. Particularly preferred Lewis acids include trimethylaluminum, triethyl aluminum, tri-isopropyl aluminum, tri-n-hexylaluminum, tri-n-octyl aluminum, dimethyl aluminum chloride, and diethylaluminum chloride.

If a single Lewis acid activator is used, it is preferably atrialkylaluminum compound, especially triethylaluminum, tri-n-butylaluminum, tri-n-hexyl aluminum, tri-n-octyl aluminum, tri n-decylaluminum, and mixtures thereof. When a mixture of two activatorcompounds is used, the compounds are desirably employed in molar ratios(activator compound 1:activator compound 2) from 6:1 to 1:1.Particularly preferred activator compounds are sequential mixtures oftriethylaluminum or tri-n-hexylaluminum (activator compound 1) withdiethylaluminum chloride (activator compound 2), or sequential mixturesof diethylaluminum chloride (activator compound 1) with triethylaluminumor tri-n-hexylaluminum (activator compound 2).

The mole ratio of total precursor activator to the electron donor in theprecursor if partial pre-activation is desired can be within the rangeof 0.1:1 to 1: 1, preferably from 0.1:1 to 0.75: 1, more preferably from0.1:1 to 0.3:1.

By the term “sequential” is meant that the second activator is notcontacted with the precursor until after contact with the firstactivator occurs, and preferably after a delay of from 10 to 60 minutes.Preactivation may be conducted in a batch process or in an in-lineprocess, and is preferably performed in an in-line fashion in which thecatalyst precursor is fully or partially activated during the period inwhich it is being conveyed to the reactor. In a preferred mode,(sometimes referred to as an in-line activation system), the precursorslurry is passed through an optional static mixer to homogenize theslurry and then past an activator injection port where activator isadded. The mixture then passes through a mixer for thoroughincorporation of activator. If a second activator is employed theprocess may be repeated until the partially or fully activated catalystmixture is injected into the reactor.

The mixers are preferably static mixers, however any suitable mixingmeans may be employed. Optionally, a vessel or a length of connectingpipe maybe provided to give an additional retention time prior toinjection into the polymerization reactor. In a desirable embodiment,the partially or fully activated catalyst precursor is passed in asubstantially plug-flow stream through any vessels, mixers andconnecting pipes or other devices in order to provide uniformlyactivated precursor composition to the reactor. This in-linemodification has the added advantage of minmizing storage time of thepartially or fully activated catalyst and allowing for direct control inreal time of the amount of activator used, resulting in improved controlof catalyst productivity. Short residence times combined with higherconcentration of reagents used in the activation results in improvedcatalyst and polymer properties, since catalyst deactivation isminimized due to the short (typically 1 minute to 6 hours) time theprecursor is in contact with the activator.

The static mixer, where employed, is preferably mounted vertically, withthe direction of flow being either up or down, to prevent solidsaccumulation in the mixer. A suitable static mixer for use hereincomprises 32 mixing elements within a 0.5 inch (12.5 mm) diameter jackethaving an overall length is 25 inches (63 cm). The static mixer elementshould be located downstream of the point where activator is injectedinto the precursor slurry. There is no requirement that the mixerelement be within a certain minimum distance of the injection point.Distances from 1 to 1000 cm are acceptable depending on the overallsystem layout and dimensions.

Static mixers function by repeatedly dividing a fluid stream passingover the mixing elements and optionally reversing the direction of flowover a small distance. Depending on the activator used, the viscosity ofthe precursor slurry, the temperature of the slurry, and other processconditions, a shorter or longer reaction period may be required foractivation of the catalyst precursor. For this purpose a suitableresidence time can be introduced into the activation process either byuse of a suitable vessel or, where plug-flow of the slurry is desired,an additional length of feed pipe or an essentially plug flow holdingvessel. A residence time zone providing an increased holding time forthe partially activated precursor can be used with both activators, foronly one activator, or for neither activator.

A preferred mode for carrying out the foregoing partial or completeactivation in-line is shown schematically in FIG. 1. In the FIGURE, theprocatalyst is introduced into a slurry feed tank 10; equipped with apump 11 for conveying the slurry to the reactor 40. The slurry passes toa first reaction zone 12, immediately downstream of an activatorinjection port 14 where the (first) activator 16, is added. Optionally,the mixture then passes to a second reaction zone 22 immediatelydownstream of a second activator injection port 24 where a secondactivator 26, may be added in a second reaction zone 22, if desired.

Each reaction zone is equipped with static mixers 20 and 28respectively. Depending on the activator compound used, some reactiontime may be required for the reaction of the activator compound with thecatalyst precursor. This is conveniently done using a residence timezone 44, which can consist either of an additional length of slurry feedpipe or an essentially plug flow holding vessel. Cocatalyst oradditional activator, is supplied from cocatalyst supply tank 46 to theslurry prior to charging to the reactor. If desired, additionalcocatalyst may be supplied to the reactor (not depicted). The reactor 40is preferably a single, continuous gas-phase reactor or continuous,dual, gas-phase reactors operating in series.

Due to the high viscosity of the slurry, poor heat transfer can resultin temperature excursion and loss of activity during activation.Depending on the catalyst precursor, degradation can start to occur attemperatures of 60° C. or higher. Accordingly, full or partialactivation according to the present invention is desirably conducted attemperatures in the range from 10 to 60° C., preferably from 30 to 45°C. Adequate mixing is desired in order to maintain relatively constanttemperatures and prevent localized catalyst decomposition due totemperature excursion.

To assure that a uniformly pre-activated catalyst precursor is suppliedto the reactor, flow through the various mixing devices and theconnecting piping should be as close to plug flow as possible. In thisregard, axial mixing and the use of residence time pots should beminimized by maintaining a high aspect ratio in the supply tubes. Apreferred IAD (length to diameter ratio) is in the range from 5 to 15.This results in a low velocity flow and minimal back mixing due tovelocity gradients in laminar flow.

After activator has been added to the catalyst precursor slurry in oneor more steps, the partially activated catalyst is then added to thepolymerization reactor where final activation by the cocatalyst occurs.Partial activation is achieved primarily by use of less thanstoichiometric amounts of the activator or by amounts that aredetermined empirically to result in incomplete activation. The remainingactivator (cocatalyst), if employed, may be added to the partiallyactivated catalyst precursor as a last step prior to entry into thereactor, or through addition to the polymerization reactor(s) or theirassociated components.

In a preferred embodiment of the invention, the final addition ofactivator occurs within 30 minutes and preferably within less than 15minutes of injection of the catalyst slurry to the reactor followed bythorough mixing and continuous plug-flow of the catalyst mixturethereafter to produce a homogeneous activated catalyst mixture. Thecomposition of the final catalyst slurry, that is the amount ofcatalyst+the amount of mineral oil diluent, is adjusted such that thefinal slurry viscosity is at least 1000 cP, preferably at least 1500 Cpas measured by a Brooldield viscometer at a shear rate of 1 sec⁻¹ at 25°C. This results in reduced catalyst settling or deposit from the slurry,especially after activation. The use of the foregoing in-line plug-flowintroduction of activated or partially activated catalyst precursor intoa reactor, especially a continuous, gas-phase polymerization reactoroperating under olefin polymerization conditions, results in uniformcatalyst properties and polymerization activity.

Complete Activation by Addition of Cocatalyst

Complete activation of the precursor by contact with activator orcocatalyst is required to achieve full activity. Suitable cocatalystsare reducing agents that are conventionally employed and known in theart, including the previously disclosed compounds used for partialactivation. Examples include hydrides, halides, and organometalderivatives of sodium, lithium, potassium, magnesium, zinc and aluminumConventionally, the cocatalysts are selected from the group comprisingaluminum trialkyls, aluminum alkyl halides, aluminum alkoxides, aluminumalkyl alkoxides, and aluminum alkoxy halides. In particular, aluminumtrialkyl- and aluminum dialkyl chloride-compounds are used. Thesecompounds are exemplified by trimethyl aluminum, triethyl aluminum,tri-isobutyl aluminum, tri-n-hexyl aluminum, dimethyl aluminum chloride,diethyl aluminum chloride, diisobutyl aluminum chloride, anddi-n-butylaluminum chloride. Butyl lithium and dibutyl magnesium areexamples of useful compounds of other metals.

Polymerization

In a single reactor configuration, the entire catalyst composition,which includes the partially activated precursor and the cocatalyst, isadded to the reactor. Alteratively, some or all of the co-catalyst maybe added to the reactor itself or to the recycle assembly comprising thereactor system In a dual reactor configuration, the reaction mixtureincluding the previously activated catalyst along with unreactedmonomers and/or the copolymer or homopolymer produced in the firstreactor, is transferred to the second reactor. Additional quantities ofpartially or fully activated catalyst and/or the same or a differentcocatalyst may be added to the reaction mixture in the second reactor orto the reaction mixture charged thereto, if desired.

The polymerization in each reactor is desirably conducted in the gasphase using a continuous fluidized bed process. A typical fluidized bedreactor can be described as follows. The bed is usually made up of thesame granular resin that is to be produced in the reactor. Thus, duringthe course of the polymerization, the bed comprises formed polymerparticles, growing polymer particles, and catalyst particles fluidizedby polymerization and modifying gaseous components introduced at a flowrate or velocity sufficient to cause the particles to separate and actas a fluid. The fluidizing gas is made up of the initial feed, make-upfeed, and cycle (recycle) gas, that is, comonomers and, if desired,modifiers and/or an inert carrier gas.

The essential parts of the reaction system are the vessel, the bed, thegas distribution plate, inlet and outlet piping, a compressor, cycle gascooler, and a product discharge system. In the vessel, above the bed,there is a velocity reduction zone, and, in the bed, a reaction zone.Both are above the gas distribution plate. A typical fluidized bedreactor is further described in U.S. Pat. No. 4,482,687, and elsewhere.

The gaseous feed streams of ethylene, other gaseous alpha-oleflns, andhydrogen, when used, are preferably fed to the reactor recycle line aswell as liquid alpha-olefins and the cocatalyst solution. Optionally,the liquid cocatalyst can be fed directly to the fluidized bed. Thepartially activated catalyst precursor is preferably injected into thefluidized bed as a mineral oil slurry. Activation is generally completedin the reactors by the addition of cocatalyst. Changing the molar ratiosof the comonomers introduced into the fluidized bed can vary the productcomposition. The product is continuously discharged in granular orparticulate form from the reactor as the bed level builds up withpolymerization. Adjusting the catalyst feed rate and/or the ethylenepartial pressures in both reactors controls the production rate.

The hydrogen:ethylene mole ratio can be adjusted to control averagemolecular weights of the polymer product. The alpha-olefins other thanethylene, if used, can be present in a total amount of up to 15 percentby weight of the copolymer and, if used, are preferably included in thecopolymer in a total amount from 0.1 to 10 percent based on totalpolymer weight. The quantity of such α-olefin can be adjusted to controlthe density of the final product.

The residence time of the mixture of reactants including gaseous andliquid reactants, catalyst, and resin in each fluidized bed can be inthe range of 1 to 12 hours and is preferably in the range of 1.5 to 5hours. Either or both of the reactors of a dual reactor system can beoperated in condensing mode, as is described in U.S. Pat. Nos.4,543,399; 4,588,790; and 5,352,749, if desired.

In a dual reactor configuration, a relatively low melt index or flowindex (or high molecular weight) copolymer is usually prepared in thefirst reactor. The mixture of polymer, unreacted monomer, and activatedcatalyst is preferably transferred from the first reactor to the secondreactor via an intercommunicating conduit using nitrogen or reactorrecycle gas as a transfer medium. Alternatively, the low molecularweight copolymer can be prepared in the first reactor and the highmolecular weight copolymer can be prepared in the second reactor.

Regardless of the reactor employed, for production of a high molecularweight product, the mole ratio of alpha-olefin to ethylene is desirablyin the range from 0.01:1 to 0.8:1, preferably from 0.02:1 to 0.35:1. Themole ratio of hydrogen to ethylene is desirably in the range of 0:1 to0.3:1, and preferably from 0.01 to 0.2:1. Preferred operatingtemperatures vary depending on the density desired, with lowertemperatures being employed for lower densities and higher temperaturesfor higher densities. Suitable operating temperature is from 70 to 110°C.

For production of a low molecular weight product, the mole ratio ofalpha-olefin to ethylene generally is in the range from 0:1 to 0.6:1,preferably from 0.001:1 to 0.42:1. The mole ratio of hydrogen toethylene can be in the range of 0:1 to 3:1, and is preferably in therange of 0.5:1 to 2.2:1. The operating temperature is generally in therange of 70 to 110° C. The operating temperature is preferably variedwith the desired density to avoid product stickiness in the reactor.

The weight ratio of polymer prepared in the high molecular weightreactor to polymer prepared in the low molecular weight reactor(referred to as “split”) desirably ranges from 30:70 to 80:20, and ispreferably in the range of 40:60 to 65:35.

The transition metal based catalyst system including the cocatalyst,ethylene, alpha-olefin, and, optionally, hydrogen are continuously fedinto the first reactor, the polymer/active catalyst mixture iscontinuously transferred from the first reactor to the second reactor;ethylene and, optionally, alpha-olefin and hydrogen, and cocatalyst arecontinuously fed to the second reactor. The final product iscontinuously removed from the second reactor. A preferred mode is totake batch quantities of product from the first reactor, and transferthese to the second reactor using the differential pressure generated bythe recycle gas compression system. A system similar to that describedin U.S. Pat. No. 4,621,952, is particularly useful in this regard.

The pressure may be the same or different in the first and secondreactors. Depending on the specific method used to transfer the reactionmixture or polymer from the first reactor to the second reactor, thesecond reactor pressure may be either higher than or somewhat lower thanthat of the first. If the second reactor pressure is lower, thispressure differential can be used to facilitate transfer of thepolymer/catalyst mixture from Reactor 1 to Reactor 2. If the secondreactor pressure is higher, the differential pressure across the cyclegas compressor may be used as the motive force to move the reactionmixture. The pressure, that is, the total pressure in the reactors, canbe in the range of 200 to 500 psig (1.5-3.6 MPa) and is preferably inthe range of 250 to 450 psig (1.8-3.2 MPa). The ethylene partialpressure in the first reactor can be in the range of 10 to 150 psig(170-1,100 kPa), and is preferably in the range of 20 to 80 psig(240-650 kPa). The ethylene partial pressure in the second reactor isset according to the amount of (co)polymer desired to be produced inthis reactor to achieve the split mentioned above. Increasing theethylene partial pressure in the first reactor leads to an increase inethylene partial pressure in the second reactor. The balance of thetotal pressure is provided by alpha-olefin other than ethylene and aninert gas such as nitrogen. Other inert hydrocarbons, such as an inducedcondensing agent, for example, isopentane or hexane, also contribute tothe overall pressure in the reactor according to their vapor pressureunder the temperature and pressure experienced in the reactor.

Desirably according to the present invention, the mole ratio ofactivator to the electron donor in the precursor employed for partialactivation in the pre-activation step (5) is within the range of 0.1:1to 1:1, preferably from 0.1:1 to 0.75:1, more preferably from 0.1:1 to0.3:1. The mole ratio of activator to the transition metal in theprecursor employed in partial activation in the pre-activation step (5)desirably is within the range of 0.25:1 to 20:1, preferably from 0.5:1to 10:1, more preferably from 0.5:1 to 5:1.

By the time of the final polymerization step, the total mole ratio ofall activator and cocatalyst employed in the present process to electrondonor is desirably in the range of 2:1 to 50:1, preferably from 3:1 to20:1, more preferably from 3:1 to 15:1. The mole ratio of totalactivator compound and cocatalyst employed in the present process totransition metal is preferably from 10:1 to 200:1, more preferably from20:1 to 100:1, most preferably from 20:1 to 50:1.

The process of the present invention unexpectedly results in a decreasein the amount of fines in the resulting polymer, in particular areduction in the level of fines particles of less than 125 μm inparticle size. The quantity of fines in the resulting product is atleast 10 percent, preferably at least 25 percent, more preferably atleast 35 percent less than the quantity of fines in a polymer producedunder the same conditions but without partial or complete pre-activationaccording to the present invention.

Although not wishing to be bound by any theory or hypothesis, theforegoing benefit of reduced polymer fines generation according to thepresent invention is believe to be due to one or more of severalpossible mechanisms:

-   1. Preactivation produces smaller catalyst particles that are    already fully activated or more easily activated upon entry into the    reactor due to their higher surface/volume ratio leading to a higher    activator/electron donor ratio or activator/titanium compound ratio.    This leads to more rapid initiation of polymerization and a longer    period of growth within the reactor leading to larger polymer    particle size.-   2. Preactivation results in higher concentration of    activator/precursor in the preactivation stage. The preactivated    particles are more uniformly advanced towards full activation,    reducing any induction period and increasing the catalyst    activity/time profile. This results in greater heat release per    particle generating faster clumping and greater structural integrity    of catalyst/polymer particles and reduced exposure of catalyst and    polymer particles to abrasion forces.-   3. Preactivation results in modification of surface of the catalyst    precursor particles causing the smaller particles to better adhere    to larger polymer particles, resulting in lower fines levels.

EXAMPLES

The skilled artisan will appreciate that the invention disclosed hereinmay be practiced in the absence of any component which has not beenspecifically disclosed. The following examples are provided as furtherillustration of the invention and are not to be construed as limiting.Unless stated to the contrary all parts and percentages are expressed ona weight basis. The term “overnight”, if used, refers to a time ofapproximately 16-18 hours, the term “room temperature”, refers to atemperature of 20-25° C., and the term “C₂H₄PP” refers to ethylenepartial pressure. In the event the name of a compound herein does notconform to the structural representation thereof, the structuralrepresentation shall control.

Test Methods

Residual Ti concentration means the titanium values in a polymer sampleexpressed in part per million (ppm), determined using X-ray Fluorescenceon a plaque prepared according ASTM D1928, Condition C. Because residualtitanium originates solely from any catalyst residue in the polymer, itis a measure of catalyst productivity. More productive catalysts resultin lower residual titanium concentrations in the polymer.

The term “Melt Index” if used herein is used interchangeably with theterm “I2” and is determined under ASTM D-1238, measured at 190° C. and2.16 kilograms and reported as grams per 10 minutes or decigrams perminute.

The term “Flow Index”, “FI” or “I21” if used herein is determinedaccording to ASTM D-1238, measured at 190° C. and 21.6 kilograms andreported as grams per 10 minutes or decigrams per minute.

The term “Melt Flow Ratio” if used herein is the ratio of Flow Index toMelt Index.

Polymer density is measured using ASTM D1928 Condition C for plaquepreparation and ASTM Method 792D for density measurement.

The terms “D10”, “D50” and “D90” as used herein indicate particularpercentiles of the log normal particle size distribution of a sampledetermined by means of a Coulter particle size analyzer using dodecanediluent and represent the particle diameter corresponding to the10^(th), 50^(th) and 90^(th) percentiles respectively of saiddistribution.

Preparation of Catalyst Precursor

A titanium trichloride catalyst precursor is prepared in anapproximately 7,500 liter glass lined vessel equipped with pressure andtemperature control, and a turbine agitator. A nitrogen atmosphere (<5ppm H₂O) is maintained at all times. Tetrahydroffuran (10,500 lbs, 4,800kg, <400 ppm H₂O) are added to the vessel. The tetrahydrofuran isrecovered from a closed cycle dryer and contained approximately 0.1percent Mg and 0.3 percent Ti. An 11 percent ThF solution oftriethylaluminum (187 lbs, 85 kg) is added to scavenge residual water.The reactor contents are heated to 40° C., and 13.7 lbs (6 kg) ofgranular magnesium metal (particle size 0.1-4 mm) is added, followed by214.5 lbs (97.3 kg) of titanium tetrachloride added over a period ofone-half hour.

The mixture is continuously agitated. The exotherm resulting from theaddition of titanium tetrachloride causes the temperature of the mixtureto rise to approximately 44° C. The temperature is then raised to 70° C.and held at that temperature for approximately four hours, then cooledto 50° C. At the end of this time, 522 pounds (238 kg) of magnesiumdichloride are added and heating initiated to raise the temperature to70° C. The mixture is held at this temperature for another five hours,then cooled to 35° C. and filtered through a 100 mesh (150 μm) filter toremove solids.

Fumed silica (CAB-O-SIL™ TS-610, manufactured by the Cabot Corporation)(811 lbs, 368 kg) is added to the above precursor solution over a periodof one hour. The mixture is stirred by means of a turbine agitatorduring this time and for 4 hours thereafter to thoroughly disperse thesilica. The temperature of the mixture is held at 40° C. throughout thisperiod and a dry nitrogen atmosphere is maintained at all times. Theresulting slurry is spray dried using an 8-foot diameter closed cyclespray dryer equipped with a rotary atomizer. The rotary atomizer isadjusted to give catalyst particles with a D50 on the order of 20-30 μm.The scrubber section of the spray dryer is maintained at approximately+5 to −5° C.

Nitrogen gas is introduced into the spray dryer at an inlet temperatureof 140 to 165° C. and is circulated at a rate of approximately 1000-1800kg/hour. The catalyst slurry is fed to the spray dryer at a temperatureof 35° C. and a rate of 65-150 kg/hour, or sufficient to yield an outletgas temperature in the range of 100-125° C. The atomization pressure ismaintained at slightly above atmospheric. The resulting catalystparticles are mixed with mineral oil (Kaydolm 350, available from WitcoCorporation) under a nitrogen atmosphere in a 400 liter glass linedvessel equipped with a turbine agitator to form a slurry containingapproximately 28 percent of the catalyst precursor.

Catalyst Precursor Partial Pre-Activation

The mineral oil slurry of precursor is partially activated by contact atroom temperature with a 30 percent mineral oil solution ofdiethylaluminum chloride (DEAC), a 50 percent mineral oil solution oftri-n-hexyl aluminum (TNA), or a sequential mixture of both activators.The catalyst precursor slurry is added to a mixing vessel at roomtemperature in an amount less than a stoichiometric amount based onLewis base present in the precursor. An appropriate amount of activatoris added while stirring. If both activators are used, the DEAC solutionis added first, and the slurry is stirred for one hour followed byaddition of the TNA solution, followed by stirring for another twohours. If only DEAC or TNA activator is used, addition is followed bystirring for at least one hour prior to use. Following partialactivation the slurry containing the partially activated precursor isretained at room temperature prior to use.

Examples 1 and 2, Comparative 1

A single, gas-phase polymerization reactor operating at 80° C. is usedto produce a high molecular weight ethylenell-hexene copolymer product.The reactor has a 14 inch (36 cm) diameter cylindrical reactor, anominal 5 to 6 foot (1.5-1.8 m) bed height, and a fluidization gasvelocity of 2 feet/sec (0.6 m/s). Comonomer content is controlled toproduce equivalent density polymers. Triethylaluminum (TEAL) cocatalystis added to the recycle gas in the form of a isopentane solution.Polymer fines are determined based on the quantity of a sample passingthrough a 120 mesh (125 μm hole size) screen. The geometric mean wasused to calculate the average particle size. Results are shown inTable 1. TABLE 1 Comp. 1* Example - 1 Example - 2 Catalyst No TNApartial TNA/DEAC preactivation activation partial activation PrecursorSize D50 23 23 23 Precursor Size D10 8 8 8 Mole Ratio Activator/THF 00.17 0.2 (TNA), 0.45(DEAC) Mole Ratio Cocatalyst/THF 5.64 6.62 6.87 C₂H₄PP psi (kPa) 38.4 (265) 35 (241) 35 (241) Residence Time [hr] 3.1 3 3.3FI (I21) [dg/min] 0.39 0.37 0.38 Density [g/cc] 0.933 0.932 0.931Comonomer (percent)** 2.25 2.4 2.6 Fines [wt %] 3.81 1.87 3.33 ResidualTi [ppm] 2.88 2.77 4.46*comparative, not an example of the invention**comonomer content in polymer

Examples 1 and 2 demonstrate reduced fines generation with respect tothe comparative polymerization. Productivity of the catalyst of Example1 as measured by residual Ti is also better than the productivity ofcomparative 1, but the productivity of Example 2 is inferior to that ofcomparative 1.

Example 3, Comparative 2

The precursor partial activation procedure of Example 1 (TNA activator,0.17 Al/THF molar ratio) is substantially repeated in combination with adual reactor gas phase polymerization process having two essentiallysimilar reactors operating in series. The cocatalyst in allpolymerizations is TEAL. Copolymer product from the first reactor ischarged to the second reactor along with additional TEAL cocatalyst andethylene monomer. Results are contained in Table 2. TABLE 2 Comparative2* Example 3 1^(st) 2^(nd) 1^(st) 2^(nd) Reactor Reactor Reactor ReactorCatalyst unactivated Partially Activated Mole Ratio Cocatalyst/ 5.896.62 6.75 7.73 THF C₂H₄ PP (psi) 32.3 (223) 100 (690) 35.6 (245) 102(703) (kPa) Residence Time [hr] 3.2 2.5 3.2 2.6 FI [dg/min] 0.7 27 0.728 Density [g/cc] 0.934 0.956 0.934 0.956 Comonomer Percent** 2.25 2.25Fines [percent] 2.34 3.11 2.06 2.64 Residual Ti [ppm] 3.14 1.58 3.251.44*Comparative, not an example of the invention**Comonomer content of polymer

The results of Example 3 again demonstrate a reduction in polymer finesin the polymer product produced in both the first and second reactors.Residual titanium values are higher after the first reactor (indicatingreduced productivity) but are reduced after completion of polymerizationin both reactors.

Examples 4, 5 and Comparative 3, 4

Another series of dual reactor mode experiments are performed using TNAas the precursor activator. In Example 5, in-line pre-activation isemployed according to the following procedure. A 2-foot (610 mm)32-element static mixer having 32 mixing elements, a 0.5 inch (12.5 mm)inside diameter, and an overall length is 25 inches (63 cm) (availablefrom Kenics Corp.) is mounted vertically with a direction of flowdownward. An activator injection point in the pipe is provided justprior to the static mixer and an up-flow, plug flow accumulator isinterposed after the mixer to provide a residence time of approximately15 to 45 minutes prior to injection into the reactor. TNA activator (50percent in mineral oil) is injected into the transfer line prior to thestatic mixer. All connections and piping are of stainless steel tubingof 0.5 inch (12.5 mm) inside diameter. The results of thepolymerizations are summarized in Table 3. TABLE 3 Comp. 3* Example 4Comp. 4* Example 5 Reactor 1st 2^(nd) 1^(st) 2^(nd) 1^(st) 2^(nd) 1^(st)2^(nd) Catalyst No activation Partial activation No activation Partialactivation C₂H₅ PP psi (kPa) 35.6 117 37.1 124 39 91 38 105 (245) (807)(256) (855) (269) (627) (262) (724) Residence Time [hr] 3 2.2 3.2 2.22.1 2.1 2.4 2.3 FI [dg/min] 0.42 27 0.43 27 0.76 16 0.7 29 Density(g/cc) 0.934 0.956 0.934 0.956 0.926 0.946 0.932 0.954 ComonomerPercent** 2.1 2.1 4.0 2.6 Fines [percent] 2.02 2.1 2.03 1.82 4.5 5.2 2.33.5 Residual Ti [ppm] 2.8 1.3 2.8 1.3 2.1 1.2 3.5 1.6 Mole RatioCocatalyst/THF 5.40 6.38 6.75 7.36 4.30 6.13 4.30 6.13*Comparative, not an example of the invention**Comonomer content of polymer

Polymer fines in the product exiting the final reactor of Examples 4 and5 decrease by 13 and 33 percent compared to Comparatives 3 and 4respectively. This result is remarkable considering the fact that theproduct formed in the 2^(nd) reactor is of a much higher density andlower molecular weight than the final product and the product of thefirst reactor and inherently more subject to fractionation and finesgeneration compared to a lower density product. In Example 5 the finalpolymer density is also higher than the final polymer density of theproduct made in comparative 4, making the reduction in fines generationeven more remarkable.

1. A gas phase olefin polymerization process comprising: (1) preparing a solution of a catalyst precursor comprising a mixture of magnesium and titanium compounds, an electron donor and a solvent; (2) adding a filler to the solution from step (1) to form a slurry; (3) spray drying the slurry from step (2) at a temperature of 100 to 140° C. to form a spray dried precursor; (4) slurrying the spray dried precursor from step (3) in mineral oil, (5) partially or fully pre-activating the catalyst precursor by contacting the slurry of step (4) with one or more Lewis Acids employing one or more in-line static mixers, and (6) transferring the partially or fully activated precursor from step (5) under plug-flow conditions into a gas phase, olefin polymerization reactor and, if the precursor is partially preactivated, adding an activator to the reactor.
 2. A gas phase olefin polymerization process comprising: (1) preparing a solution of a catalyst precursor comprising a mixture of magnesium and titanium compounds, an electron donor and a solvent; (2) adding a porous catalyst support, to the solution from step (1) to form a slurry; (3) drying the slurry from step (2) to form a solid catalyst precursor; (4) slurrying the solid precursor from step (3) in a viscous inert liquid, (5) partially or fully pre-activating the catalyst precursor by contacting the slurry of step (4) with one or more Lewis Acids employing one or more in-line static mixers, and (6) transferring the partially or fully activated precursor from step (5) under plug-flow conditions into a gas phase, olefin polymerization reactor and, if the precursor is partially preactivated, adding an activator to the reactor.
 3. The process of claim 1 or 2 wherein; 1) the catalyst precursor in step (1) corresponds to the formula: Mg_(d)(M)(OR)_(e)X_(f)(ED)_(g) wherein R is an aliphatic or aromatic hydrocarbon radical having 1 to 14 carbon atoms or COR′ wherein R′ is a aliphatic or aromatic hydrocarbon radical having 1 to 14 carbon atoms and each OR group is the same or different; M is a transition metal; X is independently chlorine, bromine or iodine; ED is an electron donor; d is 0.5 to 56; e is 0, 1, or 2; f is 2 to 116; g is >2 and up to 1.5(d)+3; and 2) the Lewis Acid of step (5) is i) one or more compounds with formula M′(R″_(n))X_((3-n)) wherein M′ is aluminum or boron; each X is independently chlorine, bromine, or iodine; each R″ is independently a saturated aliphatic hydrocarbon radical having 1 to 14 carbon atoms, provided that when M is aluminum, n is 1 to 3 and when M is boron, n is 0 to 1.5; and ii) is added in an amount such that the mole ratio of total Lewis Acid to electron donor in the precursor is from 0.10:1 to 1.0:1.
 4. The process of claim 1 or 2, wherein said Lewis Acid is; 1) one or more alklyaluminum compound(s) with formula M′(R″_(n))X_((3-n)) wherein M′ is aluminum, R″ is n-butyl, n-hexyl, n-octyl, iso-octyl, isohexyl, and n-decyl, X is Cl or Br and n is a number from 0 to 1.5; and 2) added in an amount such that the mole ratio of total Lewis Acid to electron donor in the precursor is from 0.10:1 to 0.75:1.
 5. The process of claim 4, wherein said Lewis Acid is; 1) selected from the group consisting triethylaluminum, tri-n-butyl aluminum, tri-n-hexyl aluminum, tri-n-octyl aluminum, tri n-decyl aluminum, triisoprenyl aluminum, dimethyl aluminum chloride, ethylaluminum dichloride, diethylaluminum chloride, and mixtures thereof, and 2) added in an amount such that the mole ratio of total Lewis Acid to electron donor in the precursor is from 0.10:1 to 0.30:1.
 6. The process of claim 4 wherein the slurry is partially or fully preactivated by first contacting with diethylaluminum chloride followed by tri-n-hexyl aluminum.
 7. The process of claim 1 or 2 wherein the viscosity of the slurry after addition of the activator or activators is adjusted to at least 1500 cP.
 8. The process of claim 1 or 2 in which the slurry of (2) is intimately mixed with the Lewis Acid by use of one or more vertically disposed staticmixers.
 9. The process of claim 1 or 2 wherein the one or more static mixers and connecting piping have length/diameter ratios from 5 to
 15. 10. The process of claim 1 or 2 in which said gas phase reactor is the sole olefin polymerization reactor.
 11. The process of claim 1 or 2 wherein two olefin polymerization reactors are employed. 